Process for the aromatization of dilute ethylene

ABSTRACT

Described herein are processes for the conversion of ethylene into CS+ olefins, naphthenics, and aromatics via a dual catalyst reaction utilizing a dearomatization catalyst.

INTRODUCTION

This application is a continuation of U.S. Ser. No. 15/913,503, filedMar. 6, 2018 which is a continuation of U.S. Ser. No. 14/825,815 filedAug. 13, 2015, teachings of each of which are herein incorporated byreference in their entireties.

TECHNICAL FIELD

The present disclosure relates to processes for the conversion ofethylene into CS+ olefins, naphthenics, and aromatics. The processesutilize a dehydroaromatization catalyst for the conversion of diluteethylene and other olefins, such as propylene and butenes, into apolygas material containing aromatics. The resulting polygas productproduced is of high octane and may be directly used as a gasolineblendstock or as feed to an extraction process for aromatics production,for example.

SUMMARY

Fluidized catalytic cracking (FCC) has wide applications withinrefineries for the conversion of heavy gas oils into lighterhydrocarbon. The operations intent is to produce material, which fallswithin the gasoline and diesel boiling range. A substantial quantity oflight end materials, particularly those of C3 and C4 olefins, areproduced in the FCC operation. Those light materials are often convertedinto gasoline boiling range products, using either an alkylation unit ora polygas operation.

Both of those process operations offer operational robustness to handleFCC produced feeds. The main challenge is developing a process solutionwhich can readily handle contaminants produced in the FCC, such asorganic sulphur and organic nitrogen, using heavy feeds, such as vacuumgas oil. For alkylation or polygas operations, this challenge istypically met by using relatively inexpensive catalysis.

An even lighter cut of FCC material, FCC off-gas, also presentschallenges with potential contaminants present, which can affectheterogeneous catalysts and their use. For instance, FCC off-gas streamscontain valuable ethylene and propylene in relatively small quantities,for example, in quantities less than 20%. FCC off-gas may also containheterogeneous catalyst poisons, such as hydrogen sulfide, carbonmonoxide, carbon dioxide, organic nitrogen, and water.

To address those challenges, a common and simple solution is to burn theFCC off-gas in a refinery furnace to produce useful plant heat. Othersolutions include its use as co-feed into a stream cracker forpetrochemical production or collection of its constituent compoundsusing more elaborate vapour recovery systems, involving compression,adsorption, and subsequent distillation steps. However, both cracker andvapour recovery options are typically expensive and normally onlyeconomically justified, when the FCC unit is either quite large or therefinery FCC is nearby other plants or streams which may be combined,pooled, and more economically recovered. As such, for the smallerrefiners, there is a need within the industry for a less expensiveoption, which offers higher value use for FCC off-gas, verses that ofcombustion and associated heating.

U.S. Pat. No. 3,960,978 discloses metalized (cation exchanged) zeolytes,such as ZSM-5 & ZSM-11, that comprise metals such as Zn, Cr, Pt, Pd, Ni,and Re, for example, the process technology M-Forming™ (Chen et al.,1986). The general understanding is that the ion exchange addsoligomerization capability to the aromatization functionality within thezeolite matrix that may enable the conversion of low molecular weightolefins, such as propylene, into oligomers and aromatics, via thecatalyst's dehydrocyclization functionality. Unfortunately, the U.S.refining industry has lacked wide adoption of this particulartechnology, either due to economic and/or technical reasons.

U.S. Pat. No. 4,795,844 illustrates a process for the conversion of C3and C4 olefin containing streams containing at least 50% paraffins usinga solid catalyst containing Gallium.

U.S. Pat. No. 7,419,930 shows the utility of MFI & MEL type zeolytescontaining Gallium for such conversions. U.S. Pat. No. 7,786,337provides background on the use of a dual catalyst systems containingboth zeolite and solid phosphoric acid catalysts for the production ofheavier molecules. U.S. Pat. No. 7,498,473 proposes the use of controlwater for such systems, and U.S. Pat. No. 8,716,542 illustrates the useof a dual zeolite catalyst system for handling feed streams containingsulphur species.

Other processes allow for the conversion of dilute ethylene into usefulfuels and aromatics. For instance, U.S. Pat. No. 4,899,006 providesbackground in the field of lighter olefin conversion, using operatingtemperatures between 580° C. and 750° C., over a catalyst alsocomprising zeolite with Gallium. Similarly, in 2001, Choudhary publishedresults on the aromatization of ethylene to aromatics over Galliummodified ZSM-5. (Choudhary et al., 2001).

Catalyst applications substantially involving crystalline zeolites arealso known. For example, U.S. Patent Application Publication Nos.2010/0247391, 2010/0249474, 2010/0249480, and 2014/0024870, describeprocesses using amorphous silica alumina materials, containing GroupVIII & Group VIB metals for CS+ oligomer production. One challenge hasbeen to find an economic solution, applicable for single sitefacilities, which can provide for both high conversions of both ethyleneand propylene into condensable, liquid materials. This requires both asimple process & robust catalyst solution, which can readily handleimpurities common in FCC off-gas feeds.

Accordingly, the present disclosure relates to economic processes ofconverting FCC off-gas feedstocks into naphtha boiling range componentsutilizing a robust catalyst under relatively low pressure conditions forsmall, single refinery sites. The presently disclosed processes allowfor the (1) substantial removal of basic nitrogen components from thefeedstock to protect the catalyst, (2) use of a robust catalyst whichcan handle small quantities of sulphur, and (3) economical choice of afixed bed reactor design containing at least two beds of catalyst.

DESCRIPTION OF DRAWINGS

FIG. 1 is a schematic of an exemplary process for the conversion ofmethane, ethane, and ethylene using a multibed downflow reactor.

FIG. 2 is a schematic of an exemplary process for the conversion ofethylene in high concentration within a methane and ethane containingfeed using a multibed downflow reactor and a diluent

FIG. 3 is a schematic of an exemplary process for the conversion ofethylene in high concentration within a methane and ethane containingfeed using a multibed downflow reactor and recycle compression.

DETAILED DESCRIPTION

Before the present embodiments are described, it is to be understoodthat the present disclosure is not limited to the particular processes,catalysts and systems described, as these may vary. It is also to beunderstood that the terminology used in the description is for thepurpose of describing the particular versions or embodiments only, andis not intended to limit the scope of the present disclosure.

In general, this document provides, according to certain embodiments,for processes for converting ethylene into C5+ olefins, naphthenics, andaromatics. The processes utilize a dehydroaromatization catalyst for theconversion of dilute ethylene and other olefins, such as propylene andbutenes, into a polygas material containing aromatics. The resultingpolygas product produced is of high octane and may be directly used as agasoline blendstock or as feed to an extraction process for aromaticsproduction.

In at least one embodiment, the catalyst reactors may be taken off-linefrom the processing of the FCC-off gas and regenerated periodicallyusing air and nitrogen. One catalyst reactor may be operating while theother reactor is regenerated.

In accordance with the presently disclosed process, the catalystcontains the zeolyte ZSM-5 at concentrations between 20-85% weight, Zincor Copper at concentrations less than 3% weight, and one element ofGroup IA or IIA at concentrations less than 3% weight, along with anamorphous binder comprising silica and/or alumina. The catalyst mayprovide a high conversion of ethylene in the off-gas at conditionsbetween 200-400° C., at operational pressures below 400 ps1g.

FIG. 1 depicts a process diagram for the conversion of dilute ethyleneand other light materials (such as methane and ethane) and heavierolefins (such as propylene and butylene) into larger olefinic,naphthenic, and aromatic components, and removes them from the gasstream.

Dilute gas feed containing methane, ethane, and ethylene enters theprocess as Stream 101 at relatively low temperature (120° F.) andpressure (150 psig). The gas feed contains trace nitrogen compounds (ppmlevels), which may include ammonia, amines, and/or nitriles. Thesenitrogen components are removed using a vessel containing solidadsorbent (21). The remaining stream 102 is substantially free of basicnitrogen and is further split into two streams 103 & 105.

Stream 103 is sent through exchangers and heaters prior to reaction.Cross-exchanger 24, uses the hot reactor effluent to heat the cold inletfeed gas, stream 103. Typical temperatures of approximately 300 to 600°F. are achieved using the cross-exchanger, resulting in stream 104.

Stream 104 is further heated to a reaction temperature of approximately500 to 700° F. using fired heater 22. Hot gas feed, 106, enters the topof reactor 23. It flows downward and is at least partially reacted overa 1^(st) catalyst bed containing a zeolite catalyst. As the materialreacts, it increases in temperature. Cooler feed 105 is then injectedinto the reactor, reducing the temperature prior to being introducedinto the second catalyst bed. The combined effluent, from the 1^(st)catalyst bed and injection 105, are then further reacted over a 2^(nd)catalyst bed. Hot reactor effluent, 107, exits the reactor and is cooledusing cross-exchanger 24 and cooler 25. The resulting product stream109, containing 2-phase liquid and vapour products are separated invessel 26. The majority of methane and ethane, which enter with thefeed, exit the process in Stream 110, and a substantial portion ofnaphtha boiling range material exits as Stream 111.

FIG. 2 represents a further elaboration of the invention as related toconversion of dilute ethylene. For maintaining the desired temperatures,within the illustrated multibed reactor, FIG. 1 describes the use ofcold feed injection. Here, FIG. 2 describes the use of a diluent, in theinstance where ethylene is present with a feed at a relatively highconcentration. In this case, 20% by weight of ethylene or higher, withina reactor feed stream, would be considered high and a potential diluentsolution would be of interest.

For most applications, considered herein, the unreacted material,depicted as Stream 110 in FIG. 1, is utilized as Fuel gas within arefinery complex. Often natural gas or methane is used to supplementthis Fuel gas supply need. In such an instance, rather than downstreamblending of natural gas or methane for supply, the process itself bebetter leveraged using this feed, should the feed ethylene concentrationbe considered high.

In FIG. 2, Stream 101 contains the rich ethylene gas in highconcentration, within a methane and ethane containing feed. That streamis further diluted using Natural gas make-up via Stream 100. The dilutedfeed then processed similarly to that, as described by FIG. 1.

A primary benefit to such a diluent approach, using of methane orNatural Gas, is that the design readily allows for mitigation of unitupset conditions. For this overall process, the reactions are quiteexothermic. High heat recovery via cross exchange in a case withrelatively high ethylene concentrations can lead to sudden hightemperatures within a reactor. To mitigate such temperature excursions,methane or Natural Gas dilution offers a preferred means to theinvention. Optionally, the methane or Natural Gas injection can containwater at its saturation level.

In addition to the use of Natural gas for control of the feedcomposition, an absorber (26) is depicted in FIG. 2. As the feed streamto this system is now more dilute, separation of vapour and liquid canbecome more difficult, at the operating pressures of interest.

In such a case, it is of benefit to use a heavy absorbent fluid, such asa distillate (or diesel) range material, to achieve high recovery of theheavier naphtha boiling range components within the cooled effluent gasstream 109. Stream 110 represents the heavy absorbent which is sentcounter-current to that to the absorber gas feed, 109. Inside theabsorber, 26, are sections of packing or trays to allow for efficientgas liquid contact and allow for high recovery of the C5+ materials fromthe gas stream.

Vapor recovered overhead of the absorber contains the methane present inthe feed as well as that of the injection. As mentioned, the vaporproduct may be used within a plants fuel gas header.

From FIG. 2, absorber bottoms (111) is then sent for stabilization forremoval of light ends (such as C2's & C3's) and recovery of the naphthamaterial (C5-C12) from the heavy absorbent. A portion of the heavyabsorbent can be recycled once a significant fraction of the naphtha cutis removed.

For the case where a diluent, such as Natural gas, is not available orfeasible, a process, as depicted in FIG. 3, may be used. As illustrated,vapour, 112, recovered from the absorber, which is substantiallydepleted of ethylene is utilized as the diluent for the feed to theprocess. Vapor 112 is compressed and a portion of that material isrecycled, as Stream 113, and blended with the feed. The other portion,114, is sent to Fuel gas.

These process flow diagrams are provided herein, as illustrations of thegeneral process. Certain derivations are known to thoseskilled-in-the-art, such as further integration with conventional FCClight-ends recovery equipment, various heat integration options andproduct stabilization schemes.

EXAMPLES

An example of a catalyst useful in conjunction with the illustratedprocess, is provided in Example 1. Process performance of this catalyst,under low operating pressure conditions, is provided by Examples 2, 3, 4and 5.

Example 1

200 grams of kaolin, previously calcined in air for 3 hours at 1100 C,was ground through a 60 um screen and combined with 200 grams of themolecular sieve ZSM5 possessing a Si/Al ratio of 38.2 ZSM-5, 30 grams ofsesbania powder, 120 grams silicon sol gel, 30 grams sodium silicate,and 960 grams distilled water. The combined material was mixed in a highshear twin sigma blade mixer to form a paste of a suitable consistencyto extrude through a short LID multiple 3 mm cylinders die plateextruder. The resulting extrudate was calcined in air at 843 C for 3hours. After calcination, the sodium cations are exchanged and theextrudate is calcined in air resulting in a finished catalyst.

Example 2

10 grams of 16 mesh particle size catalyst from Example 1 was loadedinto a 0.500 inch diameter 316 SS reactor tube, equipped with athermocouple, located in the middle of the catalyst bed. The reactortube was then placed in an electric tube furnace. The reactor tube washeated to 300° C. under a constant flow of research grade nitrogen,while maintaining a back pressure of 50 psig.

Once the internal catalyst bed temperature stabilized at 300° C., 5grams of liquid water was injected into the nitrogen stream at the rateof 0.25 grams/min via a feed pre-heater section.

The liquid water was completely vaporized prior to contacting the hotcatalyst bed. After steam treating the catalyst, the nitrogen feed wasdiscontinued and 3.0 weight hour space velocity of a 3.0% Hydrogen, 12%Methane, 6.5% ethane, 6.0% ethylene, 72.5% nitrogen (by volume) wasintroduced to the reactor, while maintaining a backpressure of 50 psig.The catalyst bed temperature of 300° C. was maintained for 24 hoursunder the constant hydrocarbon feed; after which, it was increased to315° C. at a rate of 1° C./minute. The entire product stream wasanalyzed by an Agilent 7890B gas chromatograph. The 12 hour averagecatalyst performance, from a time on stream of 128-140 hours, isprovided in the table below.

Ethylene Ethylene Ethylene Ethylene Ethylene Catalyst Bed Ethylene Yieldto Yield to Yield to C3 Yield to C4 Yield to CS+ Temperature WHSVConversion Methane Ethane hydrocarbon hydrocarbon hydrocarbon 315° C.3.0 92.5% 3.6% 4.6% 5.3% 14.5% 72%

The C5+ Product Composition was measured, on a mass percentage. This GCanalysis was as follows:

C5 C6 C7 Cl0 C12 isomers isomers isomers Aromatics isomers isomers 32%24% 3.5% 31.5% 4.5% 4.5%

Example 3

The reactor start-up procedure of Example #2 was repeated for Example #1Catalyst. After steaming the catalyst, the reactor temperature wasmaintained at 300° C. The nitrogen feed was discontinued and 2.6 WHSV ofa 15 mole % ethylene in nitrogen feed was introduced at a back pressureof 50 psig. These conditions were held constant for 100 hours. The tablebelow is the 10 hour average catalyst performance from a time on streamof 100-110 hours.

Ethylene Ethylene Ethylene Ethylene Ethylene Catalyst Bed Ethylene Yieldto Yield to Yield to C3 Yield to C4 Yield to C5+ Temperature WHSVConversion Methane Ethane hydrocarbon hydrocarbon hydrocarbon 300° C.2.6 98.6% 0.0% 1.4% 4.3% 12.1% 82.2%

The CS+ Product Composition on a mass percentage basis was as follows:

cs C6 C7 Cl0 C12 isomers isomers isomers Aromatics isomers isomers 15.8%15.4% 0.9% 60.5% 7.4% 0.0%

Example 4

The reactor start-up procedure of Example #2 was repeated for Example #1Catalyst. After steaming the catalyst, the reactor temperature wasincreased to 325° C. at a rate of 1° C./minute. Once the reactortemperature stabilized, the nitrogen feed was discontinued and 1.0 WHSVof a 7.0% Hydrogen, 30% Methane, 17% ethane, 15% ethylene, 31% nitrogen(by volume) was introduced to the reactor while maintaining abackpressure of 45 psig. After 150 hours at the previously statedconditions, H₂S was introduced in the feed at a rate of 200 ppm/hour.After 4 hours, the H₂S was removed from the feed. The table below liststhe 12 hour average catalyst performance pre and post H₂S addition inthe feed.

Ethylene Ethylene Ethylene Ethylene Ethylene Ethylene Yield to Yield toYield to Yield to Yield to Conversion Methane Ethane C3s C4s cs+ Pre H2SAddition 98.7% 8.4% 10.9% 4.1% 16.5% 60.3% Post H2S Addition 98.6% 6.7%8.2% 4.2% 16.5% 64.4%

As illustrated in Examples 2-4, the process yields for ethyleneconversion into C4+ materials are sufficiently high (>80%) at the lowoperating pressures of choice. For feeds containing hydrogen (Example2), a yield of greater than 86% C4+ was obtained. For feeds containingno hydrogen (Example 3), a yield of over 94% C4+ was obtained.

The difference in performance between Example 2 & 3, was identified tobe caused by the hydrotreatment activity of the metals associated withthe catalyst. The preferred catalyst metals composition, used tomaintain catalyst life, was found to allow for some hydrotreating ofethylene to ethane, which lowered the potential C4+ yield for hydrogencontaining streams.

Initially, Example 4 was made to determine how much catalyst activitywould drop, given an upstream unit upset. For FCC Off-gas operations,Hydrogen sulphide is generally removed down to low levels (typically <10ppm) so that it may be particularly used as low sulphur fuel gas.Example 4 was made to determine how much activity loss might beassociated, should a unit upset occur in an FCC amine treater, causinghigh hydrogen sulphide to enter the process.

What was surprisingly discovered, rather than direct activity loss,during the Example 4 testing on the catalyst, was that short periods ofsulphur additions were useful in improving the catalyst selectivityperformance, by selectively reducing the hydrotreating activity andsubsequently increasing the C4+ yields. This was surprising in that itoccurred without a significant loss in the catalyst'sdehydroaromatization activity.

At some point of high sulphur loading, it is rather expected thatadditional prolonged loading may cause the dehydroaromatization activityto decrease, and regeneration to be necessary in a short period of time.What example 4 illustrates is the potential use of sulphur additives todampen possible hydrotreating activity, while maintainingdehydroaromatization activity at an acceptable level, when hydrogen ispresent in a dilute ethylene feed.

Example 5

The reactor start-up procedure of Example #2 was repeated for a yzeolite modified Example #1 Catalyst. The modification of the catalystwas achieved by adding 3% wt y zeolite to the ZSM5 powder prior to themixing and extrusion step. After steaming the catalyst, the reactortemperature was increased to 345° C. at a rate of 1° C./minute. Once thereactor temperature stabilized, the nitrogen feed was discontinued and1.0 WHSV of a 20% Hydrogen, 30% Methane, 17% ethane, 15% ethylene, 18%nitrogen (by volume) was introduced to the reactor while maintaining aback pressure of 30 psig. The table below list the 24 hour averagecatalyst performance for they zeolite modified Example #1 catalyst andthe ZSM 5 standard control Experimental #1 catalyst.

Example # 1 Catalyst Ethylene Ethylene Ethylene Ethylene Ethylene GramsZeolite Ethylene Yield to Yield to Yield to Yield to Yield to CS+/literFormula Conversion Methane Ethane C3s C4s C5+ catalyst*hr ZSM5 98.1%11.2% 13.2% 8.8% 27.5% 39.8% 46.6 ZSM5 + 3% Y  100%  9..% 11.3% 7.9%27.9% 43.8% 52.3

Under high hydrogen feed content, the addition of y zeolite improves theC5+ product yield by reducing the hydro-treating activity. Y zeolitealso improves the catalyst productivity by increasing the single passethylene conversion. A 12% C5+ productivity increase (on a mass basis)was achieved with a 3 wt % addition of y zeolite to the experimentalExample #1 catalyst formulation.

LITERATURE REFERENCES

-   Chen et al., “M-forming Process”, Ind. Eng. Chem. Res., Vol. 26,    1986, pp. 706-711.-   Choudhary, V. et al., “Aromatization of dilute ethylene over    Ga-modified 7SM-5 type zeolite catalysts”, Microporous and    Mesoporous Materials. Vol. 47, 2001, pp. 253-267.

Each and every reference discussed in the present document is hereinincorporated by reference in its entirety.

What is claimed is:
 1. An exothermic process of converting ethylene in a hydrocarbon gas feed, said process comprising: a. contacting a hydrocarbon gas feed or portion thereof comprising ethylene with a nitrogen-adsorbing guard bed; b. introducing the hydrocarbon gas feed into a reactor comprising at least one heterogeneous dehydroaromatization catalyst comprising (i) ZSM-5; (ii) amorphous silica, alumina, or a combination thereof; (iii) Zn and/or Cu; and (iv) at least one exchanged metal of Group IA or IIA or lanthanide series; and c. converting ethylene and propylene in the hydrocarbon gas feed to a naphtha boiling-range product comprising at least one aromatic compound.
 2. The process of claim 1, wherein the reactor further comprises at least two beds of catalyst, wherein each bed has an independent hydrocarbon gas feedpoint.
 3. The process of claim 2, wherein said reactor further comprises a system for controlling the temperature of the hydrocarbon gas feed, wherein said system is located on at least one of the at least 2 feed points.
 4. The process of claim 1, further comprising at least one cross-exchanger to transfer heat from an effluent of the reactor or using at least one cross-exchanger to add heat to the hydrocarbon gas feed.
 5. The process of claim 1, wherein said heterogeneous catalyst comprises between about 20% and about 85% ZSM-5 by weight.
 6. The process of claim 1, wherein the heterogeneous catalyst comprises a second three-dimensional-framework, crystalline, silica/alumina species in an amount less than about 35% weight.
 7. The process of claim 1, wherein a dilution feed comprising hydrocarbons substantially free of ethylene is used to dilute the ethylene concentration in the hydrocarbon gas feed.
 8. An exothermic process for the conversion of ethylene in a C2 containing gas feed, said process comprising: a. providing at least one C2 containing gas feed, wherein said C2 containing gas feed comprises ethylene, ethane and/or methane, methanol or dimethylether; b. providing at least one heterogeneous dehydroaromatization catalyst in abed in a reactor, wherein said heterogeneous dehydroaromatization catalyst comprises (i) ZSM-5 in a range between 1-100%; (ii) amorphous silica, alumina, or a combination thereof in a range of 1-99%; (iii) Zn and/or Cu in a range between 0.05-3% weight; and (iv) at least one exchanged metal of Group IA or IIA or lanthanide series in a range of 0.05-3% weight; c. contacting said C2 containing gas feed and said heterogeneous dehydroaromatization catalyst in a bed in a reactor; d. processing, in the reactor, the hydrocarbon gas to form a naphtha boiling range product, comprising at least one aromatic component. e. optionally, stopping the C2 containing gas feed, isolating the reactor from the process, using valves, and performing a periodic regeneration of the heterogeneous catalyst to remove carbon deposits, produced from step d., using air, nitrogen, or mixtures thereof.
 9. The process of claim 8, further comprising at least two beds of catalyst, wherein each bed has an independent hydrocarbon gas feed point.
 10. The process of claim 8, wherein said reactor further comprises a system for controlling the temperature of the C2 containing gas feed, wherein said system is located on at least one of the at least 2 feed points.
 11. The process of claim 8, wherein said heterogeneous catalyst has between 20-85% ZSM-5 by weight.
 12. The process of claim 8, wherein the heterogeneous catalyst further comprises a second three-dimensional-framework, crystalline, silica/alumina species having a content of less than 35% weight.
 13. The process of claim 8, wherein a dilution feed, comprising hydrocarbons substantially free of ethylene, is used to control the temperatures within the reactor.
 14. An exothermic process for the conversion of ethylene in a hydrocarbon gas feed or portion thereof, said process comprising: a. providing at least one hydrocarbon gas feed, wherein said hydrocarbon gas feed comprises ethylene and hydrogen; b. providing at least one heterogeneous dehydroaromatization catalyst in a bed in a reactor, wherein said heterogeneous dehydroaromatization catalyst comprises (i) ZSM-5 in a range between 1-98.9%; (ii) amorphous silica, alumina, or a combination thereof in a range of 1-98.9%; (iii) Zn and/or Cu in a range between 0.05-3% weight; and (iv) at least one exchanged metal of Group IA or IIA or lanthanide series in a range of 0.05-3% weight; c. pretreating the heterogeneous dehydroaromatization catalyst to control its hydrogenation activity, by passing a gas stream containing hydrogen sulphide and or an organic sulfur compound over the catalyst, prior to introduction of the hydrocarbon gas feed; and d. processing, in the reactor, the hydrocarbon gas feed, to form a naphtha boiling-range product, said naphtha boiling-range product comprising at least one aromatic compound.
 15. The process according to claim 14, where the organic sulphur compound is a mercaptan, sulphide, or disulphide.
 16. The process according to claim 14, where the reactor comprises at least two beds of catalyst, wherein each bed has an independent hydrocarbon gas feedpoint.
 17. The process of claim 14, wherein the heterogeneous catalyst further comprises a second three-dimensional-framework, crystalline, silica/alumina species having a content of less than 35% weight.
 18. The process of claim 17, wherein the second three-dimensional-framework, crystalline, silica/alumina species is Zeolite Yor Beta zeolite. 